Co-current flow dehydrogenation system



M y 2, K. H. HACHMUTH EI'AL COCURRENT FLOW DEHYDROGENATION SYSTEM Filed Sept. 8, 1958 t FLUE GAS PAIR COMPRESSOR 44 TO SEPARATION L.-

I 4 U 2 LL! 3 a O 2 ii Ll. Ll.

I 0 2 LL! 3 O I! INVENTORS E K.H. HACHMUTH D.C. TABLER CO-CURRENTFLOW DEHYDROGENATION SYSTEM Karl H; Hachmuth' and Donald c. Tabler, Bartlesville,

Okla, assignors to Phillips Petroleum Company, a corporation of Delaware Filed Sept. s, 1958, Ser. No. 759,535 8'Claims. (11. 260-6833) heating the feed and catalyst. The present invention,

specifically directed to the dehydrogenation of aliphatic hydrocarbons, is such amethod, the heat being supplied by preheating the feed and catalyst. In such a system, accurate control of the preheating is necessary in order to supply the correct amount of heat without overheating the hydrocarbon and the catalyst. Overheating of the hydrocarbon will result in undesired thermal reactions and overheating the catalyst will greatly increase the rate of deterioration thereof.

The following are objects of our invention.

An object of our invention is to provide an improved process for the dehydrogenation of aliphatic hydrocarbons. A further object of the inventionis to provide an improved process for dehydrogenating normal butane and isopentane. A further object of our invention is to provide a co-current dehydrogenation process wherein all of the heat required is supplied by the materials charged and wherein the temperatures are suitably maintained to avoid undesirable reactions. 4

Further objects and advantages of the present invention 'will be apparent to one skilled in the art upon reading this specification, which includes:

A drawing 'showing,.in simplified form, a system utilizing the present invention.

Broadly,the invention resides in a process for dehydrogenating an aliphatic hydrocarbon containing 3 to 6, inclusive, carbon atomsin the presence of a finely divided dehydrogenation catalyst which comprises heating to a temperature 'set forth in the table depending upon the hydrocarbon present in major proportion,

heating the catalyst to a temperature at least equal to that to which the hydrocarbon is heated but not more than 150 F. and preferably not more than 100 F. above the maximum set forth for'the hydrocarbon being dehydrogenated, mixing the heated hydrocarbon and the heated catalyst, introducing the resulting mixture into one end of an elongated reaction zone at a rate sufiicient to provide a gas velocity sulficient to carry said catalyst .through said reaction zone, removing the reaction mixture and catalyst fronilthe'jsecond end of said reactio'n zone at a temperature at least 20 F. and preferably at least 50 .F. abovelthe'equilibiium temperature for the reaction involved, separating gas fi'om said catalyst, and recovering dehydrogenation products from said gas.

' From this description, it will be seen that the tem- 2,982,798 Patented May 2, 1961 5 ice being treated. The process is specifically applicable to all of the hydrocarbons containing 3 to 6, inclusive, carbon atoms and can be used to produce olefins from paraifinic hydrocarbons or for the production of diolefins from monoolefins. These compounds include, for example, propane, normal butane, 1- and Z-butenes, isopentane, normal pentane, Z-methyl-Z-butene, Z-methyl-l-butene, 1- pentene, 2-pentene, normal hexane, and 2,3-dimethylbutane. I

The drawing illustrates, in schematic form, apparatus in which the present invention can be carried out. For convenience, the dehydrogenation of butane will be described in connection with this drawing. Obviously, the other hydrocarbons are similarly treated, although the temperatures involved do vary. In this operation, butane is preheated in a furnace (not shown) to a temperature above that required for the reaction and supplied to the reactor 10 by. means of co'nduit 11. This butane is mixed with finely divided catalyst, also preheated to a temperature higher than that required for reaction, supplied by means of conduit 12. The butane picks up the catalyst and the mixture enters one end of elongated reaction chamber 10. The catalyst velocity is maintained near that of the butane flowing through the reactor by the drag of the flowing gas. The reaction proceeds rapidly because of the relatively large readily available surface area of the finely divided catalyst and the presence of the required heat at the reaction site.

At the far end of the reactor the larger part of the catalyst is separated from the gaseous products by mechanical separation means 13 and 14. In these separators the gas is removed from the center by means of conduits 1'6 and 17 while the catalyst is forced to the outside as a result of tangential force and is removed by means of conduits 18 and 19. The product stream is then quenched by means of water supplied by means of conduit 21 to a temperature below reaction conditions but still above the condensation point of the water. The cooled product then passes to secondary mechanical separation means, such as a cyclone separator 22, where all but diflicultly separable catalyst is removed from the gas. A compressor can be used between the first water quench and separator 22 when advantageous, such as in verylow pressure operation. The gas is removed from separator 22 by means of conduit '23 and the catalyst recovered by means ofconduit 24. The gas is passed to secondary quench zone 26 supplied by water from conduit 27 wherein the gas is quenched to atmospheric temperature and any residual catalyst dust removed by scrubbing with the liquid, the catalyst and liquid being removed by conduits 28. A filter can be used to remove the final amounts of catalyst. After catalyst is removed the gas is passed by conduit 29 to compressor 31 from which it is passed by conduit 32 to separation means for recovery of the desired dehydrogenation products. Separation of hydrocarbons of different degrees of saturation is Well known and is not shown in the present drawing.

The dehydrogenation reaction is favored by low pres sure, thus the preferred pressure in the reactor is from near atmospheric pressure down to as low as one tenth of an atmosphere or even lower if needed. This low pressure operation is possible because pressure drop through the reactor can be kept very low, roughly onefiftieth or the absolute pressure on the reactor itself under favorable conditions. Also the energy necessary to separate the larger part of the catalyst at the outlet of i the reactor can be furnished by the kinetic energy of the gas and catalyst without developing appreciable back pressure on the reactor.

The catalyst recovered from the reactor efliuent stream,

- and appearing in oonduits18, 1-9 and 24, is collected in perature used depend upon the particular hydrocarbon conduit 41 and'all or a portion thereogfpassed to a re- 3 generator .42 wherein the carbon is burned oif and the catalyst reactivated. This regenerator 42 is shown as the fluidized bed type well known in oil refineries where it is used to regenerate cracking catalyst. We can use the co-current flow principle of reactor for this catalyst regeneration operation. Such a system would have the advantage of low pressure drop discussed in connection with the dehydrogenation reaction. In this regenerator, air is supplied by means of conduit 43, this air serving to burn carbon from the catalyst in a fluidized bed in the regenerator. Flue gas is removed by conduit 44. Catalyst is removed from the regenera-tor by means of conduit 12 and returned to reactor 11 as set forth above.

In some situations, it is desirable to have a combination of two or more regenerators. operated in series in order to provide improved uniformity of thc'regenerated catalyst. Both of these would operate in the same method as set forth in connection with regenerator A2.

The amount of coke deposited on the catalyst during the dehydrogenation period is generally sufiicient to furnish, upon combustion with air, all the catalyst preheat needed. That is, if just sufficient air to. burn the coke on the catalyst is used to regenerate the catalyst, the reactivated catalyst leaves the regenerator at the proper temperature for return to the dehydrogenation reactor. If insufficient heat is generated from burning this coke, additional heat can be generated by adding fuel to the regenerator, such as carbon, coke, natural gas, hydrogen-rich residue gas from the process, hydrocarbon liquids or other suitable combustibles. However, because water is a temporary catalyst poison, it is preferable to avoid the common refineryhydrogen-containing fuels since use of these generally requires a preconditioning treatment of the catalyst following regeneration. Preferred is carbon or coke from an extraneous source added either as a granulated solid to the catalyst entering the regenerator, or separately to the regenerator, or burned externally in a producer" to produce a hot nitrogen plus carbon monoxide stream which is mixed with additional air to furnish a hot regeneration gas stream consisting mostly of nitrogen, oxygen and carbon dioxide.

Bypass conduit 46 communicates with conduits 41. and 12, this conduit having heating means 47 therein. By adjustment of valves 48 and 49 in conduits 41 and 46, provision is made for returning a portion of the catalyst directly to reactor 10. Preconditioning chamber 51 is provided for treatment of the catalyst following regeneration. Conduit 52 extends from conduit 12 to chamber 51 and conduit 53 extends from chamber 51 back to conduit 12. By control of valves 54 and 56, the amount of catalyst passing through chamber 51 can be controlled. Chamber 51 is provided with inlet 57 for the supply of hydrogen, mixtures of hydrogen and methane, carbon monoxide, etc. and the gas is removed through conduit 58.

An advantage of our process results from the use of the finely divided catalyst because this gives a far greater reaction rate per unit mass of the catalyst. A catalyst particle diameter of 0.01 inch (approximately 60 mesh) is preferred although the range of 40 to 200 mesh is suitable.

Dehydrogenation catalysts are well known, a particularly useful group being the oxides of the elements of the 4th, 5th, and 6th groups of the periodic table preferably supported on materials such as silicates, clays, and other refractory materials. Particularly effective is chromium oxide on alumina containing 5 to 40 percent chromium oxide calculated as Cr O For good reaction rates and satisfactory reactor volumes, the reaction should not reach too close an approach to equilibrium. Generally, a minimum of 25 to 50 F. should be provided between the exit temperature from the reactor and the equilibrium temperature conditions. The term equilibrium temperaturev conditions is defined in Industrial Chemical Calculations, 2nd

edition, by Hougen and Watson, copyright 1936, at page 460. To obtain catalyst flow sufiiciently close to the rate of gas flow, the gas velocity should be at least twice the terminal velocity of a falling catalyst particle under free fall conditions. This velocity gives a good dispersion of the catalyst in the flowing gas and permits operation with reactors having a diameter as large as 12 feet. Obviously, this gas flows depends upon the catalyst particle size.

The time of dehydrogenation is very short, being of the order of 1 to 30 seconds, preferably from 1 to 3 or 4 seconds. The amount of catalyst can be considerably less than that used in previous. dehydrogenation operations because of the large surface area available and because of the good heat transfer. Because of these advantages, we are able to operate with a catalyst concentration of 3 to 10 pounds of catalyst per pound of hydrocarbon. Preferably, the amount of catalyst is less than 5 pounds per pound of hydrocarbon.

The following example provides a disclosure of specific operating details of two runs carried out according to the process of this invention. Obviously, there can be some variation from the specific details given, although operation should be maintained within the limits set forth above.

EXAMPLE In this example butane is dehydrogenated to produce butene and a small amount of 1,3-butadiene using a chromium oxide on alumina catalyst containing about 20 weight percent chromium as Cr O Although, for economic reasons, it is desirable to operate a plant with as near as an approach to eqilibrium temperature as practical (20 to 50 F.) at the exit of the reaction chamber, circumstances, such as greater throughput rate than design, at an existing plant can make greater temperature differentials necessary. The example given illustrates this type of situation wherein the approach to equilibrium temperature is F. The details of the reactor, the flow of components, the regenerator, the flow therein and the efliuent composition are all set forth in the following table:

Table I Reactor I II Conversion, percent 30 41 Reactor pressure, atm 1. 0.26 Hydrocarbon feed rate, lbs/hr. 185,000 185. 000 Catalyst feed rate, lbs/hr 851, 000 851,000 Ratio of catalyst feed to hydrocarbon feed, lbS 4. 6 4. 6 Gas velocity at base of reactor, ftJsec 40. U 40. 0 Number of reactors 1 4 Reactor diameter, ft., each.. 5.8 5. 8 Reactor height, it., each 60 95 Pressure drop across reactor, p.s.i 0.13 O. 05 Temperature of hydrocarbon entering reactor, I 1, 1, 160 Temperature of catalyst entering reactor, F 1, 200 1, 200 Temperature of hydrocarbon+cntalyst leaving reactor, F 1, 048 996 Catalyst granule diameter, In 0. 01 0. 01 Time gas is in contact with catalyst, sec..- 1. 35 2.05 Time catalyst is in contact with gas, sec 1.51 2. 41 Overall reaction rate:

lb. butane converted/ft. catalyst surface,

hours 2.1 1.8 Regenerator:

Number of regenerator vessels. 1 1 Vessel diameter, ft., eac 19.5 19. 5 Catalyst bed depth, ft., each 23. 6 23. 6 Temperature in catalyst bed, F 1,200 1,200 Pressure above catalyst bed, p.s.i.a.. 19. 7 19. 7 Pressure drop across catalyst bed, p.s 3. 28 3. 28 Superficial gas velocity. ft./sec 1. 23 1.23 Efiiuent Composition (M01 percent):

Hydrogen 22. 8 32. 6 Methane v 1.4 1.4 Ethy 0. 4 0. 4 Ethane 0. 9 0. 9 Propylene 0. 7 0. 7 Pro n 0.2 0.2 Butndisne 1. 6 7.6 Isobutene 0. 3 0. 3 18. 8 17. 1 Normal Butane. 52. 9 38. 8

These runs show a very considerable improvement in reaction rate. over a tubular reactor currently being used and wherein the overall reaction rate is only 0.09 compared to the rates of 2.1 and 1.8 of the present invention. This improvement in conversion is obtained without the requirement for extensive heat transfer during the reaction. Furthermore the invention provides far greater utilization of the catalyst.

As many possible embodiments can be made of this invention without departing from the scope thereof, it is to be understood that all matter herein set forth is to be interpreted as illustrative and not as unduly limiting C 1200 to 1300 F. C; 1100 to 1200 F. C and C 1075 to 1175 F.

preheating said catalyst to a temperature at least equal to that to which the hydrocarbon is heated but not more than 150 F. above the maximum temperature set forth for the hydrocarbons being dehydrogenated, mixing said preheated hydrocarbon and catalyst in a ratio of 3 to 5 pounds of catalyst per pound of hydrocarbon, introducing the resulting mixture into the lower end of a vertical elongated reaction zone maintained at a pressure of 0.1 to 1 atmosphere with a gas velocity at least twice the terminal velocity of a falling catalyst particle under free fall conditions so that the catalyst is carried through said zone at a rate substantially the same as the rate of gas flow through said zone, removing the reaction mixture and catalyst from the upper end of said reaction zone at a 7 temperature at least 20 F. above the equilibrium temperature for the reaction involved, separating gas from said catalyst, and recovering dehydrogenation products from said gas.

2. The process of claim 1 wherein said hydrocarbon is normal butane.

3. The process of claim 1 wherein said hydrocarbon is a a butene.

4. The process of claim 1 wherein said hydrocarbon is isopentane.

5. The process of claim 1 wherein said hydrocarbon is normal pentane.

6. The process of claim 1 wherein said hydrocarbon is a Z-methyl-butene.

7. A process for dehydrogenating butane which comprises preheating butane to 1160 F., said temperature being higher than that required for the reaction, preheating a finely divided chromia-alumina catalyst to 1200 F., mixing said catalyst and said butane in a weight ratio of 4.6 pounds of catalyst per pound of butane, introducing the mixture into the lower end of a reaction zone maintained at a pressure of 0.1 to 1 atmosphere at a rate sutiicient to provide a gas velocity of 40 feet per second, said gas flow rate being suflicient to carry said catalyst upwardly through said reaction zone, removing the reaction mixture and catalyst from the upper end of said reaction zone at atemperature F. above equilibrium temperature and within the range of 996 and 1048 F., separating gas from said catalyst, passing at least a portion of said catalyst to a regeneration zone maintained at 1200 F. wherein carbon is burned therefrom by burning with air, returning regenerated catalyst to said mixing step, and separating dehydrogenation products from unreacted butane.

8. The process of claim 7 wherein said catalyst has a particle size of 40 to 200 mesh.

References Cited in the file of this patent UNITED STATES PATENTS 2,397,352 Hemminger Mar. '26, 1946 2,669,591 Hepp Feb. 16, 1954 2,722,476 Burnside et al. Nov. 1, 1955 2,820,072 Wood et al Jan. 14, 1958 2,889,383 Green June 2, 1959 

1. A PROCESS FOR DEHYDROGENATING AN ALIPHATIC HYDROCARBON CONTAINING 3 TO 6, INCLUSIVE, CARBON ATOMS IN THE PRESENCE OF A FINELY DIVIDED DEHYDROGENATION CATALYST WHICH COMPRISES PREHEATING SAID HYDROCARBON IN THE ABSENCE OF THE CATALYST TO A TEMPERATURE ABOVE THE TEMPERATURE AT WHICH SAID DEHYDROGENATION IS BEING CARRIED OUT, SAID PREHEAT TEMPERATURE BEING SET FORTH IN THE TABLE DEPENDING UPON THE HYDROCARBON PRESENT IN MAJOR PROPORTION 